Method and apparatus for making hydrogen-carbon monoxide mixtures



2,532,514 CARBON 3 Sheets-Sheet 1 Dec. 5, 1950 J. A. PHlNNl-:Y A

METHOD AND APPARATUS FOR MAKING HYDROGEN- uoNoxDE mx'ruREs Filed Aug.14, 1945 ln ven o r:- John .4. Phinney v f J. A. PHINNEY 2,532,514METHOD AND APPARATUS FOR MAKING HYDROGEN-CARBON MONOXIDE MIXTURES 3Sheets-Sheet 2 Dec. 5, 1-9'50 Filed Aug. 14, 1945 In Ven fon.

A i f5" .John ,4. pu'nney 7192 @y @WW-9M Dec; 5, 195o S l l J. A.PHINNEY 2,532,514 METHOD AND APPARATUS FOR MAKING HYDROGEN-CARBON vMONOXIDE MIXTURES Filed Aug.f14, 1945 3 Sheets-Sheet 3 #Hor-hey PatentedDec. 5, 1950 METHOD AND APPARATUS FOR MAKING HYDROGEN-CARBON MONOXIDEMIX- TURES John A. Phinney, Tulsa, Okla., assignor to Stanolind Oil andGas Company, Tulsa., Okla., a corporation of Delaware ApplicationAugust'14, 1945, Serial No. 610,845

4 Claims.

This invention relates to hydrocarbon synthesis and it pertains moreparticularly to an improved method and means for effecting the synthesisof hydrocarbons on a commercial scale from hydrogen and carbon monoxide.

The synthesis is exothermic and the synthol process as heretofore knownto the art has been extremely cumbersome and expensive so that it hasnot been considered economical for use in the United States. Thereforean object of my invention is to provide a simplified and improvedhydrocarbon synthesis system employing natural gas as the feed which iscommercially attractive in this country.

In the hydrocarbon synthesis reaction large amounts vof heat areliberated and it is not only necessary to remove this vast amount ofheat but it is also necessary that the conversion temperature bemaintained within relatively narrow limits. Therefore it is an object ofthis invention to provide a system wherein the synthesis temperature maybe controlled and maintained within the desired limits. v

In the preparation of hydrogen and carbon monoxide gas mixtures throughthe partial oxidation of methane with air, it is necessary to remove thesynthesis gas at high temperature, e. g., 1800 F. or higher, to obtaingas of a satisfactory composition. The reaction is not sufficientlyexothermic to attain this high outlet temperature without preheating thereactants to a temperature of between about 1000 and 1200 F. Preheatingthe reactants to a higher temperature has additional advantages since itpermits the utilization of carbon dioxide made in the synthesis toproduce additional quantities of carbon monoxide by the endothermicreaction Various systems have heretofore been proposed for the reactionof methane with air or oxygen and for the exchange of heat between hotprod- Y ucts and cold reactants. The prior art systems,

however, have a number of disadvantages. For

the gases pass into the main body of the chamber. In some othermodifications cold oxygen or air is charged to a burner where combustiontakes place with a preheated gas. If the flame is extinguished for anyreason the possibility arises for mixtures of unreacted oxygen andhydrocarbon gas to fill the vessel with the result that an explosionoccurs when the mixture comes in contact with a hot surface.

Therefore, another object of the invention is to provide an improvedsystem for converting hydrocarbon gas such as natural gas containingmethane into mixtures of hydrogen and carbon monoxide suitable for usein the synthesis of hydrocarbons. These and other objects will becomeapparent as the detailed description of this invention proceeds.

In general, the objects of this invention are attained by separatelypreheating air, natural gas and recycle carbon dioxide by heat exchangewith hot synthesis gas, combining the preheated gases and-reacting themover a catalyst at about 1900 F. Prior to the preheating by heatexchange with the product gases, the air, hydrocarbons and carbondioxide are heated to about 850 F. being increased to a temperature ofabout 1625 F. in the heat exchange tubes. The gaseous product mixtureleaves the exchanger at about 1000 F. and is supplied to a hydrocarbonsynthesis zone after further cooling to about 400 F. The hydrocarbonsynthesis is conducted in the presence of a finely divided catalystmaintained in a dense turbulent suspended phase and cooling tubes areprovided for the extraction of the exothermic heat of reaction.

The catalyst for synthesis reaction can be either of the cobalt type orof the iron type. The cobalt type promotes the reaction:

2:cHz+CO (CHz) I-i-:nHzO (3) and the iron type catalyst promotes thereaction:

3xHz+3CO 2 (CH2) I-i-.rHzO-i-xCOz (4) In either case the catalyst shouldbe in finely divided form so that it can be fluidized by gases or vaporsflowing Iupwardly through the catalyst at low velocity. The use ofcatalyst particles of such structure, shape, and size as to be uidizedby upflowing gases of the dened velocities is an important feature ofthe invention.

The temperature of the synthesis step when employing an iron typecatalyst usually is within a range of between about 450 and about 675F.; for example, about 550 F.

With a cobalt type catalyst the temperature of the synthesis step isusually within the range of about 225 and 450 F., for example, betweenabout 325 and 395? F. The lower temperatures tend toward the productionof heavier hydrocarbons such as waxes and the higher temperatures tendtoward the production of lighter hydrocarbons such as gases with theoptimum liquid yield being obtained within the preferred range of 325 to395 F.

The preferred raw material for the production of carbon monoxide andhydrogen is natural gas, but carbon monoxide-hydrogen mixtures can beprepared from other sources. Some natural gases such as those obtainedfrom the Hugoton field are of low sulfur content and need not bedesulfurized. When necessary, however, the natural gas is first freed ofhydrogen sulfide and organic sulfur compounds by conventionaldesulfurization processes.

As pointed out above, a considerable amount of heat must be supplied forthe gas reforming operation and this heat is preferably obtained by heatexchange with the hot gas reforming products. The hot gases are passedthrough a catalyst bed. The catalyst is preferably a group VIII metal ormetal oxide which can be either unsupported or supported on clay,kieselguhr, silica gel, alumina, Super Filtrol, and the like. Thecatalyst can be promoted by a metal or metal compound, for example, theoxide of aluminum, magnesium, uranium, chromium, molybdenum, vanadium,and the like.

The space velocity through the gas reforming catalyst should besufficient to give a contact time of between about 0.1 and about 45, forexample, between about 0.4 to about seconds. The temperature of thisoperation is preferably at between about 1500 and about 2100 F., forexample, 1900 F. and the outlet pressure may be as high as about 250pounds per square inch, for example, about 225 pounds per square inch.This reforming operation converts the methanecarbon dioxide-air mixtureinto a gas consisting chiefly of hydrogen, carbon monoxide, and diluentnitrogen, this gas mixture being hereinafter referred to as synthesisfeed gas.

The synthesis feed gas is cooled and introduced into the synthesisreactor where it is exothermically reacted in the presence of a finelydivided catalyst. In order to prevent the temperature level of thereactor from gradually increasing, i. e., to remove the heat ofexothermic reaction, a plurality of individually controlled bayonet typecooling tubes may be extended into the mass of turbulent catalyst.

In systems of this type catalyst solids of small particle size areuidized by upflowing gasiform materials within the reactor so that thecatalyst within the reactor is maintained in a turbulent liquid-likedense phase, the extreme turbulence of theA suspended catalyst particlesserving to maintain substantially the entire mass of catalyst at auniform temperature. The catalyst particles are of the order of 2 to 200microns or larger, preferably to 100 microns in particle size. Withvertical gasiform fluid velocities of the order of about 0.5 to 5,preferably between about 1 and 4, for example, about 2 feet per second,a liquid-like dense phase of catalyst is obtained in which the bulkdensity is between about 30 and about 90 per cent, preferably betweenabout 40 and about 80, e. g., about 60 per cent of the density of thesettled catalyst material. The vertical velocity of the gasiform fluidsis in any event regulated so as to produce a turbulent suspension ofcatalyst within the reactor.

An active iron type catalyst can be prepared by a number of methods wellknown in the art and can, for example, be of the precipitated typesupported on Super Filtrol or other carriers. Alternatively an ironcatalyst of the type used for ammonia synthesis can be employed, suchcatalyst ordinarily being prepared by oxidizing iron in a stream ofoxygen, fusing the oxide and crushing. The catalyst can be reducedbefore use, preferably with hydrogen, at a temperature of between about600 and 1500 F. If desired make-up catalyst can be supplied to thereactor as the oxide, the oxide undergoing reduction within the reactor.Various promoters such as alkali metal compounds may be added. Anothermethod employs the decomposition of iron carbonyl to form an iron powderwhich may be sintered and ground. Catalyst particles without a supportmay have a bulk density as high as about to 150 pounds per cubic foot,whereas the bulk density of iron catalyst supported on Super Filtrol orsimilar carrier may be as low as about 10 pounds per cubic foot.

It is also contemplated that a cobalt type catalyst can be usedconsisting essentially of supported cobalt metal either with or withoutone or more promoters such as oxides of aluminum, cerium, magnesium,manganese, thorium, titanium, uranium, zinc, zirconium, and the like.The cobalt support is preferably an acid-treated bentonite or clay suchas Super Filtrol or other material of low calcium or iron content. Othersupports include materials such as koalin, alumina, silica, magnesia,and the like. The cobalt-to-carrier ratio may be varied between about 5to 1 and 0.1 to 1. The catalyst may be reduced before use, preferablywith hydrogen at a temperature of between about 350 and about 500 F.Likewise the reduction can be carried out within the reactor proper withthe synthesis gas as the reducing medium.

Instead of the cobalt or iron catalyst, I may employ catalyst of thenickel type or the ruthenium type. The above catalysts are all known tothe art and inasmuch as no invention is claimed in their composition ormethod of preparation, further description is not believed necessary.

The invention will be more clearly understood from the followingdetailed description read in conjunction with the accompanying drawingswhich form a. part of the specification and wherein:

Figure 1 is a schematic flow diagram of a complete process design forpracticing the invention;

Figure 2 is an enlarged View of the feed gas preparation unit andFigures 3 and 4 are sections taken along the lines 3--3 and l--Irespectively;

Figure 5 is a diagrammatic representation of another modification of agas reforming apparatus; and

Figure 6 is a partial section along the line B--S in Figure 5.

In the specific example a. system will be described for handling naturalgas which consists essentially of methane. The application of theinvention to other charging stocks and to plants of different capacitywill -be apparent to those skilled in the art.

The charging stock is introduced via line III, passed through productheat exchanger Il and introduced into the reformer I2 at a temperatureof about 800 F. and a pressure of about 240 pounds per square inch. Insome instances it will be necessary to desulfurize the natural gasbefore it is supplied to the system and in such an event anyconventional desulfurization system can be used. Air is supplied Vialine I3 to the compressor I4 which may be in two or more stages. The airis Withdrawn from compressor I4 at a temperature of about 390 F. under apressure of about 250 pounds per square inch. This high pressure air issupplied via line I3 to the reformer I2 after passing' through productheat exchanger I5 wherein the temperature is increased to a'bout 850 F.Recycle carbon dioxide is supplied at a pressure of about 240 pounds andat a temperature of about 380 F. via line I6. This substantially purecarbon dioxide is obtained by means of absorbing the CO2 from a portionof a recycle gas fraction returned to the hydrocarbon synthesis step.

The reformer I2 is provided with a header I6, the gases being introducedinto separate tubes I1 wherein the gases supplied by lines I0, I3, andI5 are preheated. separately to at least 1500 F. The gases thuspreheated pass downwardly from the preheat tubes I1 at high velocity andenter a reactionspace I8. The lower ends of the tubes I1 pass through acatalyst bed I9 supported intermediate the ends of reformer I2. When theair and methane are preheated separately to at least l500 F. theair-methane mixtures 0f widely varying composition are spontaneouslyinflammable and when the preheated gases enter the reaction space I8 theoxygen is immediately consumed. The partially reacted gases from thereaction vzone I8 are thoroughly mixed and enter the bed of catalyst I9where the conversion of residual methane, carbon dioxide, and watervapor to hydrogen and carbon monoxide is effected. A suitable catalyticmaterial for this purpose is nickel on high temperature fire brick orthe like. The product gas, leaving thecatalyst bed I9, then passescountercurrent to the incoming gases within tubes I1. A temperature ofabout 1900 F. is maintained within the catalyst bed I9 and heat isextracted by the incoming gases in tubes I1 to attain a temperature ofabove l500 F., for example, about 1620 F. in the tubes I1. The productgases then leave the chamber I2 via line 20 at a temperature of about1000 F. and under a pressure of about 225 pounds per square inch.

Natural gas in being heated to about 1500" F. deposits carbon on thewalls of the preheating Vtubes I1 and if the carbon is allowed toaccumulate it would render the unit inoperable. To

y avoid the continued deposition of coke, natural gas and air arealternately passed through the tube, the air removing the carbon ascarbon dioxide. The reformer I2 includes three separate preheatingsections I1, each section handling alternately hydrocarbons, carbondioxide and air, and carbon dioxide in sequence to avoid mixing any airand hydrocarbon gas in any preheating section. The flow through theselected preheating tubes I1 may be controlled by an automatic timingdevice (not shown). Although banks of tubes are shown for each section,it is contemplated that a single tube can be used in each section inplace of the banks of tubes.

The product gases in line from reformer I2 are split and a portionpassed via line 2I through heat exchanger II and another portion throughheat exchanger I5. The product gases are thence cooled, commingled withrecycle gas from line 22 and introduced into the synthesis reactor 6,23. A uidized bed of synthesis catalyst such as nely divided iron ismaintained within this reactor 23 which also contains bayonet tubes 24'to remove the heat of reaction by the generation of steam.

The synthesis gas stream in line 20 is introduced into the reactor 23 ata low point therein, preferably through a distributor plate. The reactor23 comprises an elongated vessel having an enlarged upper section 29wherein the catalyst settles out from the reaction products by reason ofthe reduced velocity therein. The velocity reduction results from thereduction in volume of the reactant gases and from the increasedcross-sectional area of the reactor 23. The reaction conditions and theproperties of the catalyst phase within the reactor 23 may besubstantially the same as described in general terms above. Cooling iseffected within the reactor 23 by a number of bayonet or thimble tubes24, the cooling tubes 24 extending through this enlarged settling zone29 and into the dense catalyst phase.

Boiler feed Water is introduced via line 30 at about F. under a pressureof about 400 pounds per square inch. The feed water passes through theheat exchangers 3| and 32 and is introduced into the steam drum 33 atabout 450 F. and at about 420 pounds per square inch. 400 pounds steamis Vwithdrawn via line 34 from the steam drum 33. t

Cooling water is supplied by line 35 to the manifold section 36 of thecooling system, said manifold communicating with the inner tubes of thethimble assemblies 24. If desired the manifold 36 may be divided into aplurality of zones whereby separate banks of tubes 24 can beindependently controlled. The -heat exchange medium passes downwardlywithin the inner tube and flows upwardly within the outer tube into themanifold 31. The relatively hot heat exchange fluid which in the case ofwater will be converted into steam is withdrawn from the manifold 31 vialines 38 into steam drum 33. Condensed water is recycled via line 35from the steam drum 33 to the manifold 36.

Reaction products from which a substantial portion of the catalyst hasbeen removed are withdrawn through line 25 and introduced into anexternal catalyst separation means 26. This system may comprise aplurality of stages of cyclones, ceramic catalyst filters, or the likeand it should be understood that the means 26 can be mounted within thereactor 23, for example, within the space provided by the enlargedsettling zone 29. A substantially catalyst-free reactant ellluent isremoved by conduit 28 for further processing. It is preferred totransfer the recovered catalyst via line 21 or its equivalent into thedense phase turbulent suspended catalyst maintained within the reactor23. If desired the recovered catalyst may be withdrawn from the systemvia line 21a,for regeneration and the like before recycle, but normallythe catalyst is maintained within the system for extended pe` riods oftime.

The gasiform product in line 28 is cooled in heat exchanger 32 andintroduced into wax knock-out drum 4U. The drum 40 can be operated at apressure of about pounds per square inch and hydrocarbon productsboiling above 400 F. are removed via line 42. The products boiling below400 F. are withdrawn from the wax separator 40 via line 4I, passed inheat exchange with boiler feed water in exchanger 3| and charged in thevapor state to a scrubber tower 43. The vapors enter gas scrubber 43 ata temperature of about 275 F. and the scrubbed gases are removedtherefrom by line 44 at a temperature of about 240 F. Within thescrubber 43 partial condensation takes place and the product liquid andgas are scrubbed by a recycle water stream to remove traces of catalystsand oxygenated compounds, the scrubber water being introduced into thetower 43 via line 45. The liquid products and condensed Water enter theseparator drum 46 which is provided with bailles 41 and 46 whereby themixed water and hydrocarbons are separated to concentrate the catalystparticles which are withdrawn as a catalyst water slurry via line 50 andreturned into the reactor 23. The water fraction is withdrawn via line45 and recycled by pump 52 to the scrubber 43. The net water separatedin drum 46, containing the oxygenated compounds, is withdrawn via lineand is removed for further treatment as desired.

The gases from the scrubber 43 are passed vis. line 44 through condenser52 where Water is removed from the gases for recycle via line 53 andline 45 to the scrubber 43. The gases from condenser 52 are removed at atemperature of about 100 F. and 185 pounds per square inch via line 54into water knock-out or dry drum 55. In the dry drum 55 residual wateris removed and recycled via line 56. Following the mist extraction indrum 55, all or part of the gases are passed by line 51 into a carbondioxide absorber 58. The lean absorber medium is introduced to absorber58 via line 58. The absorber is operated at a slightly elevatedtemperature, say 13o-150 F. to avoid hydrocarbon condensation. TheCO2-rich absorber medium is withdrawn from the absorber 58 via line 60and supplied to the carbon dioxide stripper 6|. Carbon dioxide isremoved from stripper 6| via line 62 at a temperature of about 230 F.and a pressure of about 55 pounds per square inch. The eilluent iscooled in cooler 63 and introduced into separator drum 64. The gaseouscarbon dioxide is removed by line I5 from the drum 64 at about 100 F.and about 50 pounds per square inch. The liquid from drum 64 is pumpedthrough line 65 into the stripper 6|. The carbon dioxide in line |6 iscompressed by means of compressor 66 to about 240 pounds per square inchwhereby it is heated to about 380 F. This high pressure carbon dioxideis recycled to the reformer |2 by line I6.

The hydrocarbon liquid fraction recovered from separator drum 46 iscarried by line 61 through heater and supplied to the heavy productsstripper 16. The liquid product is introduced at about 350 F. and thestripper 16 is operated at about 10 pounds per square inch. Steam issupplied to the stripper by line 11. The bottoms from the stripper 16are withdrawn via line 18 to storage and comprise a synthol gas oil.

The gases in line 51 which are not subjected to carbon dioxide recoveryare either recycled to the reactor 23 via line 22 or sent to the productrecovery system. The portion which is recycled to the reactor 23 isblended with the CO2-free gases recovered via line 80 from the CO2absorber 58. These recycled gases are compressed by compressor 8| andintroduced into the reactor feed line at about 220 pounds per squareinch.

The product gases in lines 51 and 83 are introduced into the oilabsorber 84 and scrubbed with an absorber oil introduced through line 85for the recovery of condensable hydrocarbons. In the oil absorptionsystem, 90 per cent recovery of net pentanes is attained. The unabsorbedgases which leave the top of the absorber 84 through line 66 aresupplied to the drying drums 81.

The rich absorber oil is withdrawn from absorber 84 via line 80 andpasses through heat exchanger 8| to rich oil still 82 provided withreboiler 83. The lighter products stripped from the gas oil fraction inthe heavy products stripper 16 are passed via line 84 into the rich oilstill 82 at a. high point therein. Stripping steam can be introducedinto rich oil still 82 via line 82a and the overhead from the still 82passes via line 85 through cooler 86 into separator-reflux drum 81. Thedrum 81 is adapted to effect separation between hydrocarbons and Waterand the water fraction is removed via line 88. This fraction is rich inoxygenated compounds which can be recovered as useful products. A refluxfraction is passed via line 88 and pump |00 into the rich oil still 82.The net hydrocarbons withdrawn overhead from the still 82 ultimately areintroduced into the stabilizer |0|.

The product liquid is fractionated to slightly above 400 F.Eli/*gasoline in the rich oil still 82, the net gas oil b 'ing removedas a part of the lean oil via line |0/'from the tower bottom. Thebottoms from the still 82 are passed in heat exchange with the rich oilin line 80 and pumped to the top of the absorber 84 via line 85. The netgas oil product can be withdrawn via line |03 with the stripped heavyproduct in line 18 from the stripper 16.

The raw synthol gasoline and light ends are compressed to 235 p. s. i.and stabilized, the overhead product of propanes and butanes beingrecycled to the absorber. actually a debutanization, the highlyunsaturated butane fraction being desirable as a charge to thepolymerization section.

The liquid hydrocarbons recovered in drum 81 are withdrawn via line |04and passed via pump |05 into the stabilizer |0| together withhydrocarbons from knock-out drum ||2. -The gases from drum 81 arecompressed by compressor ||0, passed through cooler and introduced intoknock-out drum ||2. The uncondensed fraction is compressed by compressor||3 and raised to a pressure of about 250 pounds per square inch and atemperature of about 295 F. This fraction is introduced into thestabilizer |0| via line ||4. A

, stabilized gasoline fraction is withdrawn from stabilizer |0| via line||5 at a temperature of about 370 F. and a pressure of about 235 poundsper square inch. A reboiler 6 is provided on stabilizer |0|. Theoverhead from stabilizer |0| is withdrawn at a temperature of about 200F. and a pressure of about 240 pounds per square inch via line ||1through cooler ||8 into reux drum ||8. The reflux liquid is withdrawnfrom the bottom of ||8 and passed by pump |20 and line |2| into thestabilizer |0|. The stabilizer gases are removed from reflux drum ||8 byline |22 and recycled to the oil absorber 84 where the uncondensablegases are purged from the absorber-stabilizer system and passed via line86 to the bauxite drums 81.

The stabilized hydrocarbon fraction in line ||5, which comprisesapproximately equal proportions of Cs, Cs and C1+ hydrocarbons, arevaporized in furnace |23 and introduced by line |24 into clay-treatingdrums |25. The clay- 76 treated fractions are introduced into reruntower This operation is |26 via lines |21 and |28. A reboiler I3Isupplies heat to the rerun tower. Steam is introduced at a low point inthe rerun tower |26 by line |30 whereby the polymer fraction is strippedand a polymer traction is withdrawn as bottoms from rerun tower |26 vialine |28. A ilnished synthol gasoline fraction is withdrawn overheadfrom rerun tower |26, cooled in |32 and introduced into accumulator drum|33. -A portion of theproduct liquid can be passed via. line |34 andpump |35 into the rerun tower |26 as reflux, whereas the net productionof nished synthol gasoline is withdrawn by line |36. Steam condensate iswithdrawn from drum |33 by line |31.

, The off gases inline 86'from the oil absorber 84 contain the'net C4and lighter hydrocarbons.

These gases are dried by passage through a bauxite system comprisingdrums 81. The clay in drums 81 is periodically and alternatelyregenerated by passing hot residual gases from the sulfuric acidabsorber through the bauxite drums in a direction countercurrent tothe'no-rmal ow of gases. The dry gases are` charged to an acid absorber|40 by line |45 where the unsaturated light hydrocarbons are absorbed inrecirculated and fresh 98% sulfuric acid supplied by line |4I. Theabsorber |40 is operated at about 250 pounds per square inch and at atemperature of about 100 F. The lean gases are withdrawn from the acidabsorber |40 via line |42 and can be utilized as fuel or heated in |43and used as a regeneration gas for the bauxite drums `81,`the'regeneration gases and moisture being sent by line |44 to fuel.

The rich acid from the acid absorber |40 is transferred via line |46 andpump |41 through heat exchanger |48 and heater |49 wherein thetemperature is increased from about 100 F. to about 200 F. The heatedrich acid in line |50 is then introduced into the upper portion of thepolymerization reactor or soaking drum |5I which isv maintained under apressure of about 550 pounds per square inch and a temperature oi' about200 F. Under the conditions within the soaking chamber |5I -thebutylenes are converted almost completely while propylenes are convertedonly partially.

The reaction products stream withdrawn via line |52 from soaking drumI5I 'is split with one portion being recycled via line |53, pump |54,and line to the soaking drum I5| and the other portion passing throughthe heat exchanger |48 and cooler |55 whereby the temperature of theproduct stream is reduced from about 200 F. to about 100 F. The cooledproduct stream in line |52 is passed in series through the acid settler|56 and the polymer settler |51, line |58 being provided to eiect thetransfer between the two settlers Some of the acid from the acid settler|56 is returned via lines |59 and |60, together with fresh ma'. ,--upacid, to the acid absorber |40 via line I4I. Another portion of theseparated acid from settler |56 is recirculated viav line |62, pump |63,and lines |46 and |50 to soaking drum |5I. Spent acid is withdrawn fromthe system via line |6|.

In the polymer settler |51 a heavy product can be withdrawn via line |64and another fraction of the polymer product can be Withdrawn via line|65 and introduced into debutanizer |66. Prior to introduction into thedebutanizer |66 the hydrocarbon fraction can be caustic-treated by theintroduction of a 2% by weight caustic solution into the line |65 byline |61.

The debutanizer |66 is provided with a reboiler |68 and can be operatedat a pressure of about 250 poundsper square inch with a vapor outlettemperatur of about 125 F. The overhead in line |14 from the debutanizer|63 comprising substantial amounts of propylene is cooled in cooler |68to a temperature of about 110 F., liquids being accumulated in drum |10.Condensate water can be Vwithdrawn Via line |1I and the liquefiedhydrocarbons via line |12. A portion of the hydrocarbon condensate isreturned via line |13 to the debutanizer |66 as reflux and the remainderis recycled to the polymerization reactor |5| by lines |12, |46 and |50.

The total polymer is removed from the debutanizer |66 as bottoms and isintroduced into caustic wash tank |16 via line |11 at a tempera ture ofabout 350 F. and under a pressure ci about 250 pounds per square inch.In the caustic wash tank |16 the hydrocarbons are treated with hotdilute caustic to break up any residual acid ester. It is contemplatedthat the caustic solution can be recycled by pump |18 and line |18 andthat the spent caustic can be removed from the wash drum |16 by line|80.

The caustic-washed polymer is introduced into rerun tower I8 via line|82 where a motor gasoline fraction nd a fuel oil cut are recovered. Thererun toyr can be operated with an outlet temperature/ I about 180 F.under a pressure of about iive pounds per square inch and a bottomtemperature of about 425 F. The overhead in line |83 from rerun towerI8I is cooled in cooler |84 to a temperature of about 115 F. and a motorgasoline fraction accumulated in accumulator |85., Steam is introducedinto the rerun tower |8| at a low point via line |86 and steam conden-'sate separated in the accumulator |85 is withdrawn via line |81. Aportion of the gasoline fraction from accumulator |85 can be returnedvia line |88 near the top of the rerun tower I8| as reux and the netproduction of polymer gasoline withdrawn from the system via pump |88and line |90 through cooler ISI at a temperature of about F. The fueloil cut is recovered as bottoms from the rerun tower |8| at atemperature of about 425 F., is cooled by cooler |92 to about 100 F. andwithdrawn from the system via line |93.

Another modication of the reformer apparatus is illustrated in Figures2, 3 and 4 of the drawings.

The reformer comprises an outer cylindrical shell 200 having an internalinsulation of refractive material 20| such as fire-brick. The top of thechamber 200 is provided with a divided header 202, each of baffles orpartitions 203 restricting the introduction of a selected gas to a givenbank of tubes 204, 205, or 206. Separate inlets 201, 208, and 209 supplythe gases to the header 202. The tubes 204 and 206 are supported byupper tube sheet 2I0 and lower annular tube sheet 2| I. The lower tubesheet 2|| is supported on angleiron 2I2 with asbestos gasketing 2I4between the sheet 2II and angle-iron 2I2 to allow for downward expansionof the tubes along bolts 2 I3, the angle-iron 2 I2 being shaped to thecontour of the chamber 200. An inner vessel 2I5 is adapted to retain thecatalyst and is xed at its upper end to the annular tube sheet 2| I.Stub spacers 2I6 are carried by the inner vessel 2I5. The inner vessel2I5 is corrugated at 2I1 to allow for expansion or contraction. It willbe noted that the inner vessel or canister 2I5 is xed to the lower tubesheet 2 I I and that the expansion of the tubes 204 and 206 over thebolts 2I3 is transmitted to the corrugated portion 2|1 of the innervessel 2I5.

The lower end of the inner vessel 215 is provided with an alloy screen218 over a refractory brick grid 219. This brick grid 219 extendsupwardly into the inner shell 215 for about six inches, the brick gridand the inner vessel 215- being supported by brick arches 220. Thesearches are in the primary reaction zone 221 and the hot combustion gasespass around these arches before entering the inner vessel 215 throughgrid 219 and screen 218. Figure 4, which is a view taken along 4--4 inFigure 2, illustrates the details of this arrangement.

The reformer illustrated in Figure 2 operates in essentially the samemanner as that illustrated in Figure 1. The gases are introduced via theinlets 201, 208, and 209 and are preheated separately in thecorresponding tubes to a temperature of at least 1500 F. The thuspreheated gases pass downwardly from the preheat tubes around the innervessel 215 and into the reaction space 221. The partially reacted gasesfrom the r'eaction zone 22| are thoroughly mixed in passing through thearches 220, grid 219, and screen 218 and enter the inner vessel 215which retains a bed of catalyst 222. In passing through the catalystbed, the conversion of residual methane, carbon dioxide, and water vaporto hydrogen and carbon monoxide is effected. A suitable catalyticmaterial for this duty can be any one or more of those known in theprior art, for example, nickel on a refractory support.

The product gases leave the inner vessel 215 and pass upwardly throughthe preheat zone. 223 countercurrently to the incoming gases. Atelnperature of about 1900 F. is maintained within the catalyst bed 222and heat is extracted from the reaction products by the incoming gasesin tubes 204, 205 and 206 to attain` a preheat temperature above about1500 F., for example, about 1625 F. If desired, disc and doughnutbailles 224 can be provided within the preheat zone 223. The productgases, in effecting this preheat of feed gases, leave the chamber 200via line 225 at a temperature of about 1000 F. and under a pressure ofabove about 200 pounds per square inch.

A suitable method for bringing such a reformer on stream includescharging cold air and gas to burners 226 in the reaction zone 221 untilthe catalyst bed 222 and the interior of the chamber have been heated tothe desired operating temperature. Initial temperature control can beobtained by admitting additional methane through the preheating coils ortubes and as soon as a temperature above about 1500 F. is attained, thenormal operation described above can be followed.

Figure 5 diagrammatically illustrates a further embodiment of a reformerfurnace which can be lemployed in my system. The apparatus illus- 235,236, and 231 to preheat zones 238, 239 and F 240 which may comprisetubes or bundles of tubes within the preheat zone 232 in a, lowerportion of the chamber 230. The catalyst bed 234 is retained on a screenor perforated member 241, the arches 243 supporting the assembly withinthe '12 chamber 230. Alternatively a water-cooled metal grid can be usedin the place of screen 24| and the arches 243.

The preheated gases by-pass the catalystl chamber 234 and are commlngledwithin the initial reaction zone 233. If desired a plurality of mixingdevices can be employed and this is particularly desirable when each ofthe preheat zones 238, 239 and 240 comprises a bundle of tubes. Theinitial reaction products pass downwardly through the catalyst bed 234,through the screen or grid 24| and into the preheat zone 232 wherein theincoming gases within the preheat zones 233, 239, and 240 areheat-exchanged with the product gases in a manner similar to thatdescribed in connection with the other modifications of the reformerfurnace. The relatively cooled reaction products are withdrawn from a,low point in the reformer 230 through ports 244.

From the above detailed description it will be apparent that the objectsof this invention have been accomplished and that a vastly improvedsystem for synthesizing normally liquid hydrocarbons from carbonmonoxide and hydrogen derived from natural gas has been provided.Although the flow diagram has been described with reference to a singlereformer and a single synthesis reactor, it is contemplated that aplurality of such reactors and gas reformers can be used in series orparallel. For example, in one process design involving the principle ofthis invention two reformers and three hydrocarbon synthesis reactorswere employed.

Although certain preferred embodiments of apparatus and operatingconditions have been described to illustrate the invention, it should beheating separately conned streams of methane,

air, and carbon dioxide and avoiding coke deposits in the separatelyconned preheating zones by passing successively therethrough, methane,carbon dioxide,'air and carbon dioxide whereby coke produced in thepreheating of methane is oxidized by the air and each zone is purged bycarbon dioxide before and after the introduction of air.

2. Apparatus for the production of mixtures of hydrogen and carbonmonoxide from natural gas which comprises a vertically elongatedreaction chamber, a headerv at one end of said chamber, partitionswithin said header dividing it into at least three sections, separateinlets for each of said sections, a bank of tubes communieating witheach of said header sections, an upper tube sheet supporting said tubesand closing one end of said chamber, a lower annular tube sheet xed tothe lower end of said tubes, an expansion joint means for supportingsaid lower tube sheet, an open-ended, substantially cylindrical vesselwithin said chamber and spaced from the inner Wall thereof, the upperend of said vessel being fixed to said annular tube sheet, a screenacross the-now area of the lower end of said cylindrical vessel, a gridmeans supporting said Screen, a perforated arch support means for saidgrid and cylindrical vessel said perforated means being spaced from thatend of the chamber which is remote from said header, and a gas outletmeans from said chamber remote from said annular tube sheet.

3. In a method for making synthesis gas from methane wherein a gaseousmixture of hydrogen and carbon monoxide is produced by the controlledoxidation of gaseous hydrocarbons, the improvement which comprisespreheating sepa,- rately conned streams of gaseous hydrocarbons, of air,and of carbon dioxide to a temperature above about 1500" F., comminglingthe separately preheated streams to eiect combustion of a portion of thehydrocarbon gases tol produce water vapor and additional carbon dioxide,reacting the residual hydrocarbon gases with carbon dioxide and watervapor in the presence of a reforming catalyst at a temperature withinthe range of between about 1800-2200 F. whereby substantial proportionsof hydrogen and carbon monoxide are produced, and avoiding thedeposition of coke in the separately coniined preheating zones bysuccessively passing therethrough separate streams of gaseoushydrocarbons, of car- -bon dioxide, of air and of carbon dioxide,whereby any coke deposit produced in the preheating of the hydrocarbongases is oxidized by the air and each zone is purged by carbon dioxidebetween the air and' hydrocarbon gas preheating steps.

4. An apparatus for the two-stage conversion of natural gas to producemixture of hydrogen and carbon monoxide which comprises a verticallyelongated shell, a header in the upper part of said shell, partitionswithin said header dividing it into a plurality of sections, inlets forsupplying a separate stream of reactant gas to each section of saidheader, an annular tube-sheet intermediate the ends of said shell,separate tubes within said shell communicating with each section of saidheader and extending downwardly to said tube-sheet for discharging saidgases into an annular space below said tube-sheet, an openy endedcanister having upper and side walls for retaining a quantity ofcatalyst intermediate the ends of said shell below said tube-sheet, theupper walls of the canister extending to said tube-sheet, the side wallsof the canister being spaced from the side walls of the shell to formsaid annular space. and the bottom of the canister being spaced from thebottom of the shell to form a reaction space, said reaction spacecommunicating with said annular space and said canister to provide apath of flow for gases discharged from said tubes whereby gases fromsaid tubes ow thru said annular space, then thru said reaction space,then upwardly through said quantity of catalyst and finally around andin heat exchange relationship with said tubes, and an outlet forwithdrawing gases from the upper part of the shell.

JOHN A. PHINNEY.

REFERENCES CITED The following references are of record in the le ofthis patent:

UNITED STATES PATENTS Number Name Date 1,843,063 Burke Jan. 26, 19322,051,363 Beekley Aug. 18, 1936 2,185,989 Roberts Jan. 2, 1940 2,210,257Pyzel et al Aug. 6, 1940 2,243,869 Keith June '3, 1941 2,258,511Leprestre Oct. 7, 1941 2,274,064 Howard Feb. 24, 1942 2,319,508Leprestre et al. May 18, 1943 2,324,172 Parkhurst July 13, 19432,347,682 Gunness May 2, 1944 FOREIGN PATENTS Number Country Date288,662 Great Britain- Apr. 16. 1928

1. IN THE REFORMING OF METHANE WITH AIR AND CARBON DIOXIDE TO PRODUCEHYDROGEN AND CARBON MONOXIDE, THE IMPROVEMENT WHICH COMPRISES PREHEATINGSEPARATELY CONFINED STREAMS OF METHANE, AIR, AND CARBON DIOXIDE ANDAVOIDING COKE DEPOSITS IN THE SEPARATELY CONFINED PREHEATING ZONES BYPASSING SUCCESSIVLEY THERETHROUGH,